Production of hydrocarbons from heavy hydrocarbonaceous residues by two stage processwith the use of inert solids



1957 E. D. BOSTON 2,813,916

PRODUCTION OF HYDROCARBONS FROM HEAVY HYDROCARBONACEOUS RESIDUES BY TWQ STAGE PROCESS WITH THE USE OF INERT soups Filed Nov. 20, 1953 2 Sheets-Sheet 1 '50005/ BURNER r .STR/PPERX 40 I W PRIMARY {600D} 42 PRODUCTS J OUEIVCH 26 4| 25 TRANSFER 44 iii 4ooE VAPOR/25R 2 SECONDARY 36 2o 7 (POOR) 1 43 IOOOF mooucrs 1 1 i l 4 I 32 l I500. 2| m H .0 29 IL 6 2 L H '30 "20F RES/Z7.

h 23 FEEDI 23 STEAM I300F STEAM IMF 2 DRIER STEAM FIGURE"/ E a'wafa' 0. Bosfon lnvenfor B zZw' hm Attorney Nov. 19, 1957 E. D. BOSTON 2,313,915

PRODUCTION OF HYDROCARBONS FROM HEAVY HYDROCARBONACEOUS RESIDUES BY TWO STAGE PROCESS WITH THE USE OF INERT souns Filed Nov. 20, 1955 2 Sheets-Sheet? FLUE- "CYCLONE 400 I36 STRIPPER I35 PRIMARY (6000M I38 PRODUCTS STEAM OUENCH |20oE I32 I06 SECONDARY {POOR} PRODUCTS I05 l I DRIER a no .1 E LJ STRIPPEl-P w I08 f I sTEAM I03 WEE/0110M STEAM STEAM FEED STEAM FIGURE '2 Edward 0. Boston Inventor By )7), Wf/orney United States Patent PRODUCTION OF HYDROCARBONS FROM HEAVY HYDROCARBONACEOUS RESIDUES BY TWO PROCESS WITH THE USE OF INERT Edward D. Boston, Westfield, N. J., assignor to Esso Research and Engineering Company, a corporation of Delaware Application November 20, 1953, Serial No. 393,375

11 Claims. (Cl. 260--666) This invention relates to the production of petrochemicals such as olefins, diolefinsand aromatics by high temperature fluid coking of heavy hydrocarbon oils. In particular it relates to a multi-stage coking process wherein the vapor product from a low-temperature stage serves as the feed of a high-temperature stage so that the petrochemicals can be recovered from the latter stage in relatively high concentrations.

High temperature fluidized coking of heavy hydrocarbons such as reduced crude or petroleum residuum is assuming increasing importance as a process for making valuable. chemical raw materials from a relatively cheap feed stock. However, in schemes heretoforeproposed for this purpose comparatively low yields of distillate were obtained. Furthermore, the gaseous product contained the desired ethylene fraction in relatively low concentration, due to dilution with hydrogen and methane. In fact, the major expense involved in such an operation is incurred with product recovery rather than in the conversion itself.

It is, therefore, the object of the present invention to produce olefins, diolefins and aromatics such as ethylene, propylene, butylenes, butadiene-1,3, isoprene, cyclopentadiene, benzene, toluene, Xylene, styrene, methylstyrene, indene, methyl indene and the like. A more specific object is to produce these petrochemicals so that they can be recovered at the least possible cost, that is, in the largest possible yield and with a minimum of dilution by hydrogen and methane. A still more specific object is to provide a high temperature fluid coking process and apparatus wherein the heavy hydrocarbon feed is cracked in stages at progressively higher temperatures in such a fashion that the very refractory and high boiling feed fractions are kept out of the principal cracking zone and thus are prevented from diluting the desired product with hydrogen and methane which tend to be produced in large amounts when the high boiling fractions are pyrolized. These and other objects, as well as the general scope and specific nature of the invention will become more clearly apparent from the subsequent description, particularly when read in connection with the accompanying drawing.

In the drawing Fig. l is a schematic illustration of a specific embodiment of the invention wherein residuum is partially converted by thermal cracking in a first stage in the presence of finely divided inert solids at a relatively low temperature between about 950 and 1100" F., the resulting vapor product is separated from the first-stage solids, and converted into desired unsaturated and/or aromatic hydrocarbons by contact with other relatively hot solids at a temperature of about 1300 to 1600 F., while conversion of the more refractory feed portions which are adsorbed on the separated first-stage solids is completed separately from the first-stage vapors.

Fig. 2 is a schematic illustration of another embodiment wherein the residuum feed and the hot. contact solids pass through a series of stages in essentially countercurrent flow. with respect to each other.

2,813,916 Patented Nov. 19, 1957 According to the invention it is possible to accomplish etficient conversion of relatively cheap stocks such as vacuum residuum, long residuum, reduced crude, whole crude, or heavy catalytic cycle stocks and the like into valuable olefins, diolefins, aromatic hydrocarbons and other chemical raw materials. Generally the feed stocks are such that a major or at least a substantial fraction thereof cannot be completely vaporized at atmospheric pressure without thermal decomposition. The feeds may have a gravity between about l0 and 20 API, a Conradson carbon content of about 5 to 40 weight percent, and a boiling range extending well above 950 or even above 1300 F. Typical of such stocks are residua, having an initial boiling point above about 1050 F., and gravity of about 11 API.

The multi-stage conversion of the present invention is usually carried out at substantially atmospheric pressure, e. g. at 0 to 25 p. s. i. g., though substantially higher pressures as well as subatmospheric pressures may be used when special considerations warrant this. The temperature in the first stage preferably is kept between about 950 to 1100 F., after which the resulting vapors are separated and further converted in one or more consecutive stages at temperatures which may range from about 1300 to 1600 F. Depending on reaction temperature, refractoriness of the. feed, desired degree of conversion, and other specific conditions the vapor residence time in the first stage is kept down-to about. 0.3 to 8 seconds, preferably about 0.6 second at 1100 F. or about 4 seconds at 950 F. The vapor residence time in the one or more consecutive stages may add up to as much as a total of about 1.5 seconds at 1300 F. down to extremely short times such as about 0.01 second at 1600 F. Preferred conditions may involve residence times of about 0.8 sec- 0nd at 1300 F. to about 0.05 second at 1500 F.

The wet solids separated from the low-temperature stage and containing the more refractory, unconverted high boiling fractions of the feed are dried by coking in a separate zone, preferably in the form of a dense fluidized bed. Here the residence. time may be considerably longer, e. g. 5 seconds to 5 minutes, and the temperature may be anywhere from about 950 to 1600 F., and preferably at least 50 to F. higher than the first conversion stage, e. g. 1200 to 1500' F. The pressure again preferably is substantial atmospheric. In general, of course, the conditions in this drying zone are quite flexible since the resulting hydrogen-rich vapors are of relatively low value anyway and are recovered separately from the primary'reaction products.

The solids from the drier may then be stripped and a portion thereof supplied directly'to the first conversion stage to serve as contact solids-therein. Alternatively, it is possible to pass essentially allof the solids from the drier through a burning zone where the solids areheated to about l350 to 1700 F. by partial combustion and the heated solids are then supplied to thevarious conversion stages as needed; Since the process-normally produces more coke than is consumed in the combustion stage, net coke product maybe withdrawn at any convenient point. Also, since the solid particles used in the process tend to grow in size by continuous deposition of coke, at least a portion of the circulating solids is preferably passed through a grinding zone such as aball mill or through a high velocity gas jet grinder so as. to maintain the particle size of the solids essentially constant;

Thecontact solids used in the system'may, be any'material which is substantially catalytically inert and refractory such as coke, sand or pumice. Coke is usually used in preference to all others. at temperatures below 1350 F., though above this temperature sand may be preferred in view of the excessive water-gas reaction which may occur between stripping. steam and the hot coke. The size of the particles may range predominantly between about 40 and 500 microns, with the greatest concentration of particles preferably in the 80 to 300 micron range. The solids pass through the conversion stages or transfer line reactors while suspended in the hydrocarbon vapors and, where necessary, an auxiliary inert gas such as steam may also be added so as to maintain the gas velocity in the desired range. This gas velocity in the several transfer line reactors may range between about 5 and 100 ft./ sec. The suspension of solids in gases passing through the transfer line reactors typically will have a density of about 0.5 to 25 1bs./cu. ft. The weight ratio of solids to hydrocarbons may be between about 2 and 30 in the first contacting stage, and about 8 to 70 in subsequent stages.

Referring to Figure l of the drawing, a specific embodiment of the invention will now be described.

Fresh feed such as South Louisiana vacuum residuum having a gravity of 11 API, a Conradson carbon content of 17 weight percent and an initial boiling point of about 1050 F., preferably after preheating to about 700 F., is fed into the vaporizer or first conversion stage through line 1. The vaporizer 10 may be a transfer line about 1 to 5 feet, e. g. 3 feet, in diameter and of such length as to give a vapor residence time of about 1 to 3 seconds, e. g. 2 seconds. Since the vapor velocity is preferably held between about 10 to 50 feet/ second, e. g. at feet/ second, a practical reactor may have a length of about 3 to 80 feet long, e. g. 50 feet. Also present in reactor 10 are hot finely divided solids which are introduced into the bottom of the reactor through line 2 at a weight ratio of about 2 to parts, preferably 8 parts of solid per part of hydrocarbon feed. These incoming solids will be at a temperature of about 1200 to 1400 F., e. g. 1300 F. Finally, it is usually also necessary to introduce an extraneous gas such as superheated steam through line 3, at a rate sufficient to maintain the gas velocity in reactor 10 within the aforesaid limits so as to make it possible for the solids and the hydrocarbon feed to pass through the reactor in the form of a suspension having a density of about 1 to lbs/cu. ft., e. g. 2 lbs./cu. ft.

The temperature in reactor 10 near the liquid feed inlet may be about 1050 to 1200 F., e. g. 1120 F., and somewhat lower near the exit end of the reactor, due to the endothermic nature of the evaporation and mild cracking taking place in the reactor. For instance, if the mixture of liquid hydrocarbon feed and hot solids is 1120 F. near the inlet of the reactor, the temperature near the exit end of the reactor may be about 1000 F. While the hydrocarbon feed passes through reactor 10, it is partially vaporized and converted, giving about 85 to 98 weight percent, e. g. 95%, of hydrocarbon vapors boiling below about 1050 F. The unvaporized refractory high boiling remainder of the feed remains adsorbed on the solids, at least partially in liquid form.

After passage through reactor 10 the mixture of vapors and solids is rapidly separated in a dust separating device such as a conventional small cyclone 4 or in a spiral separator of the type described hereafter with reference to Figure 2. It is desirable that this separation be of relatively short duration compared to the vapor residence time in reactor 10, or in any event the total time elapsed from the time feed is injected into reactor 10 to the instant that the solids are substantially separated therefrom in separator 4 should not exceed the residence time limits indicated earlier herein, e. g. 0.5 second at an average temeprature of 1050 F. Otherwise the refractory high boiling portion of the feed will begin to decompose excessively and give off undue amounts of hydrogen and methane. On the other hand, the completeness of the separation between vapors and solids is not especially important, since it is suflicient that the bulk of the wet solids, e. g. at least 90% of the total solids present, be removed from the vapors before the solids are passed through dip leg 6 to a drier or stripping vessel 30, the operation of which will be described later.

The separated completely vaporized feed portion is passed from separator 4 to a second reactor 20 through line 21. Reactor 20 again preferably is an elongated conduit or transfer line about 1 to 5 feet in diameter and about 3 to 50 feet long. For instance, it may be a tube about 2 feet wide and 20 feet long. Upon admission into reactor 20 the hydrocarbon vapors are mixed with a stream of hot solids introduced through pipe 29. These solids will be at a relatively high temperature, e. g. at about 1500" F., and may be added to the hydrocarbon vapors in a weight ratio of about 15 parts of solids per part of hydrocarbon. The vapors in the resulting suspension thus will be rapidly heated to approximately 1500 F. As the suspension passes through transfer line reactor 20, the vapors will be extensively cracked and somewhat cooled in the process. The exit temperature thus may be about F. lower than the solids inlet temperature while the vapor residence time in reactor 20 may be about 0.2 second, the pressure being about 10 p. s. i. g. The vapor velocity in reactor 20 may be kept at about 100 ft./sec., some of this velocity being due to the steam which is added through aeration lines 23 for the purpose of carrying the hot solids into the reactor. The cracked vapors are separated from the solids in a separator such as cyclone 24, and the separated vapors are finally quenched to a temperature below about 600 F. by injection of cooling medium such as water or steam or recycled liquid hydrocarbon fraction or even cool solid particles through line 25. The separated product vapors may then be drawn off through line 26 and passed to a conventional recovery system which may comprise a primary fractionator for separating the tar, gas oil and lighter distillate fractions; a compressor for the overhead 0 gas; and suitable recovery equipment for the light gases such as oil absorption or low temperature distillation.

Table I below shows a typical product distribution obtainable by following the inventive procedure just described, as compared with a product composition obtained under the same conditions when the pitch feed is injected directly into a transfer line coking reactor in accordance with prior art practice.

TABLE I Product yield and dzstrzbutlon Typical Comparable Product of Composition, Invention w. Feed Di- (Wt. percent rect to Ooker based on (Wt. percent feed) based on feed) Hydrogen 0. 2 O. 4 4. 3 0.9 1 12. 4 2 10.6 2. 7 2. 7 9. 7 8. 7 Propane 0.4 0.4

03 and lighter.... 29. 7 29. 7 Butadlene 1-3. 2. 7 2. 4 Butylenes..-.- 4. 8 4. 1 Butanes 0. 2 O. 6

Total 0 7. 7 7. 1 Isoprene 0. 8 0. G Piperylene 1. 0 0. 9 Oyclopentadiene 0. 7 0. 0 Other Cs. 1.6 2.0

Total C5 4. 1 4. l 3. l 2. 2 l5. 3 l3. 1 25. l 26. 8 15.0 17.0

1 49 mol percent in C2 and lighter cut.

2 34 mol percent in C2 and lighter cut.

In comparing the results of the two runs summarized in Table I, several important differences are noted. First of all, by operating in accordance with the invention the total ethylene yield is increasedby about and that of propylene by more than 10%. Furthermore, itis even more significant to observe that according to the invention the concentration of ethylene in the C2 and lighter cut is about 63 weight percent or 49 mol percent, as against about 52 weight percent or only about 34 mol percent in the prior art. This increase in ethylene concentration is of course very important as far as recovery of the. olefin isconcerned.

In addition it is. seen that the. invention produces appreciably increased amounts of all other olefins, notably the butylenes, and also of aromatics as indicated by an almost 50% greater yield of benzene. On the other hand, the formation of the less desirable products such as tar and coke isrepressed.

The solids recovered in separator 24 may be passed through dip leg, 27' either to the same stripper and drier 30 as that mentioned previously in. connection with the first-stage solids, or; as an alternative, separate drying units may be used for each stage. In vessel 3f the solid are preferably kept in the form of a dense fluidized bed by injecting stripping steam. through line 31 upwardly through the vessel at a rate sufficient to give a superficial gas velocity of about 0.5 to 5 ft./sec. within the dense bed, e. g. 2 ft./ sec. Here the conversion of the refractory high boiling desidues adhering to the solids is substantially completed, with the production of relatively low molecular weight vapors and a dry, solid carbonaceous residue. For this purpose the solids in vessel 30 are preferably kept ata temperature intermediate between the respective temperatures prevailing in transfer lines 10 and Zil, e; g. at about 1300? F. Thus, the heat content of the solids fed into vessel 30 from separators 4 and 24 is normally'sufficient to maintain vessel 30 at the desired temperature. However, if a higher stripping temperature should be preferred, further heat may be added to vessel 30 by introducing into it hot solids directly from. the burner, or by other appropriate heating means. The relatively. poor vapor products. which will contain at least about 30 mol percent of hydrogen and methane, are finally recovered through line 32. Normally, liquid fractions may be condensed out of this stream by cooling, whilev the uncondensed gas may be used as fuel.

The hot stripped and. dried solids may be withdrawn from vessel 30. andpassed to a burner 40 through standpipe 35 and riser 36. A portion of the hot solids, equal to about one-third of the solids going to burner 40, may be drawn oif fromriser 36 substantially at or near stripper temperature andpassedthrough line 2 to supply heatto the first-stage vaporizer 10 as previously described. As normally the. temperature of the solids coming from drier 30will be, suflicient to satisfy the heat requirements of the vaporizer 10, an inert gas such as steam will be injected through line 37 at the foot of riser 36 to serve as lift gas therein. However, if hotter solids are preferred in vaporizer 10, an appropriate amount of an oxygen-containing gas such as air may be injected through line 3'7 toproduce some combustion in riser 36 before the solids are drawn off through line 2. Such combustion in riser 36 has the advantage over combustion directly in drier 30in that it doesnot dilute the hydrocarbon vapors liberated in the drier. Of course, where such combustion is carried out in riser 3.6, it will be desirable to inject a stripping gas through line 38 into the solids in line 2; so as to keep the flue gases from diluting the valuable hydrocarbons evolved in' vaporizer 10.

The remainder of the solids present in line36 are passed to aheating zone 40 where they are reheated in any known manner to a temperature at least slightly above that desired in the-high-temperature conversion stage 20, e. g. to 1500 This heater 40. may be either in the form of another transfer line reactor as shown, air being injected through line 41, or the heating may be accomplished by burning a-portion of they circulating solids while '6 in the form of a dense bed. Still further alternatives. may include heating the solids by contact with hot flue gases, or by indirect heat exchange, all of which are well known.

Referring to illustrated transfer line burner 40, the nnburnt portion of the solids is separated from the flue gases by a dust separator such as cyclone 47 and the hot separated solids, preferably after stripping with steam from line 44 in vessel 42, are finally passed through standpipe 43 and riser 29 back to the main conversion stage 20 to supply the required heat of cracking as previously described.

It will be noted that the main advantage of the invention is due both to keeping the refractory hydrogenproducing bottoms out of the principal reaction zone, and to the fact that the cracking reaction is kept more nearly homogeneous throughout the main conversion stage inasmuch as the feed to this stage is essentially entirely gasiform and therefore permits use of high vapor velocities and correspondingly high turbulence throughout. In contrast, if the liquid residuum feed were injected in conventional manner directly into the bottom of dilute phase transfer line coker 20, not only would the refractory feed portions produce large volumes of hydrogen and methane, but even the lighter feed portions would become unduly degraded by a combination of undercracking and over-cracking portions thereof. The latter would be due to the fact that the initial vapor velocity would have to be kept relatively low so as to allow for a considerable increase in velocity as more vapors are produced by the progressive cracking of the liquid feed. Consequently, such cracking of an at least partially liquid feed would not be nearly as homogeneous as in the case of the all-vapor feed present in the cracking zone of the instant invention.

An alternative embodiment of the invention will nowbe described in which the hydrocarbons to be cracked pass through a series of stages of increasing temperature in countercurrent contact with the heat-carrying contact solids, with separation of solids and vapors between stages.

Referring specifically to Figure 2, preheated residuum feed of the type previously described is injected through line 101 into the bottom of the first conversion stage or vaporizer 110. This stage is again a transfer line through which the hydrocarbon feed passes at a gas velocity of about 10 to 50 ft./sec., together with about 15 parts by weight of finely divided coke per part of hydrocarbon. This coke may be fed into the stage 110 from line 102 with the aid of an inert carrier gas such as steam injected at 103. Additional steam may also be injected through line 104 to help atomize the liquid feed and to bring the gas velocity in the vaporizer 110 to the required value. The temperature of the coke in line 102 may be about 1200 F. As a result of the introduction of these hot solids into vaporizer 116, the latter may be close to about 1000 F. at its vapor outlet as the feed is partially vaporized and cracked.

The resulting vapors are rapidly separated from their burden of solids which contain the unvaporized refractory portion of the feed. Preferably the residence time of the feed in low temperature vaporizer 110 is again kept within the aforementioned ranges, depending on temperature, nature of feed, etc. For instance, the hydrocarbons may remain in contact with the vaporizer solids for about 2 seconds at 1000" F. Separation of the solids from the resulting vaporscan be accomplished in a variety of wellknown ways, a convenient way being simply to bend the end of the vaporizer line 110 in the shape of a loop 106'. The resulting change in direction will cause the suspended coke particles to bethrown centrifugally against the outer wall: so thata substantially solids-free vapor stream can bewithdrawn from the inner wall of the loop through line 105;

After this separation the wet solids may be passed through dippipe 107 into a drier :and stripper vessel 108 essentially similar to vessel 30 described in connection with Figure 1. In this vessel the wet residue adhering to the coke particles is essentially dried by completing the coking reaction so as to convert all liquid hydrocarbons into vapors and a solid carbonaceous residue. This, of course, requires a considerable residence time, c. g. the average residence time of the coke particles in fluidized vessel 108 may be about sec. to 5 minutes. The resulting vapors, which will tend to be high in hydrogen and methane, may be withdrawn through line 109 for use as fuel or the like. The dried solids may be withdrawn from vessel 108 through line 112, for passage to a heater or burner. This heater may be essentially of the same type as burner 40 described in connection with Figure l, or any equivalent device. The purpose of such a heater is to raise the heat content of the solids sufficiently so as to heat the hereafter described highest temperature conversion stage 130 to the required temperature, e. g. to 1400 or 1500 F.

The separated first-stage vapors are substantially free of any liquid feed portions and of solid particles, and comprise predominantly an essentially uncracked gas oil fraction boiling up to about 1050 F., as well as substantial amounts of thermally cracked naphtha and gas. These vapors are withdrawn from the aforementioned vaporizer 110 through line 105 and passed to the inlet of the next conversion stage 120, where they become admixed with the solids present in the U-bend 121. These solids are at a temperature of about 1300 or 1400 F. inasmuch as they have been separated from the next higher temperature stage. As the vapors from line 105 enter the transfer line reactor 120, they will form a suspension with the solids from U-bend 121 and the resulting suspension will pass through reactor 120 at a vapor velocity of about 50 ft./sec. The suspension may contain the hot solids in a weight ratio of about 16 parts of solids per part of hydrocarbon, and the suspension may have an apparent density of about 2 lbs./cu. ft. As the suspension passes through reactor 120, the gas oil type vapors are extensively cracked into the desired low molecular weight compounds, usually with an attendant drop in temperature. Thus the temperature in reactor 120 may range from about 1350 or 1400 F. at its inlet to about 1200" F. at its outlet after passage through reactor 120 the cracked vapors are separated from the solids which may be done substantially in the same manner as already described in connection with the vaporizer stage 110. After separation the solids from the second conversion stage 120 are allowed to drop through line 102 so as to be available for introduction to the inlet of the first conversion sage 110, as already described.

The substantially solid-free vapors withdrawn from the second conversion stage through line 125 may then be passed to still another conversion stage 130, to be cracked therein at still higher temperature so as to provide the very highest temperature cracking zone immediately adjacent to the product quench. Such a three-stage system makes it possible to operate under conditions approaching the ideal which requires having a progressively higher temperature as the cracking reaction proceeds.

However, even a two-stage system is similarly feasible and offers major yield advantages over a single-stage operation.

The hydrocarbon vapors from line 125 may be mixed in the high temperature reactor 130 with hot coke coming at about 1500 F. through line 142 from the heater previously mentioned. The re-cracked vapors will finally be separated from the hot solids in loop 136. The separated solids, cooled to say about 1400 F. while passing through reactor 130, will be passed through line 132 and U-bend 121 to the aforementioned intermediate conversion stage 120 to supply the heat of reaction thereto. To prevent subsequent degradation the separated vapors are promptly quenched by injection of a cooling medium such a water, steam or cold solids. t rough line 138 and finally recovered from line 135. Quenching by indirect heat exchange can serve the same purpose. The composition of these product vapors will be substantially similar to those described in connection with the two-stage process of Figure 1, except that the yield of distillate fractions will be somewhat greater and the yield of tar less by a corresponding amount.

While all stages have been shown as operating with upward flow, it is possible to operate alternate stages with opposite directions of flow, thereby saving on total reactor height.

Wherever ratios or percentages are given in this specification and the claims, they are to be taken as referring to a weight basis unless specifically stated otherwise.

Having described the general nature of the invention as well as two specific embodiments thereof, it will be apparent to those skilled in the art that within the scope of the appended claims the invention may be embodied in still other modifications without departing from the scope or spirit hereof.

The claims:

1. A process for converting heavy oils containing constituents nonvaporizable without cracking which comprises the steps of vaporizing a heavy oil in an elongated vaporization zone by passing said heavy oil and heated particulate solids there through at a temperature between about 950 and 1100 F., separating tarry solids containing adhering liquid hydrocarbons from the vapors so formed, contacting said vapors with additional amounts of particulate solids at a chemicals coking temperature and passing the resulting mixture through an elongated conversion zone, separating from the effluent of said elongated conversion zone gasiform conversion products and spent solids, passing said spent solids and said tarry solids to a fluid coking zone and completely drying said solids therein, recovering gases comprising hydrogen and methane from said fluid coking zone, passing a portion of the dried solids from said fluid coking zone to said elongated vaporization zone, heating the remainder of the dried solids to above said chemicals coking temperature and transferring the solids so heated to said elongated conversion zone.

2. In a process for converting heavy liquid hydrocarbon oil containing at least a substantial fraction boiling above about 900 F., into olefinic and aromatic hydrocarbons, the improvement which comprises mixing said liquid hydrocarbon oil with finely divided inert solid particles passing the resulting mixture in the form of a gasiform suspension through a confined elongated vaporizing zone at a temperature between about 950 and 1100 F. so that the oil is partially vaporized, promptly separating the resulting vapors from the unvaporized oil which adheres to the solid particles, mixing the separated vapors with dry finely divided inert solid particles, passing the resulting mixture in the form of a gasiform suspension through a confined elongated conversion zone at a temperature between about 1300 and 1600 F., recovering cracked product vapors rich in olefins, diolefins and aromatics from said conversion zone, passing said separated wet solids from said vaporizing zone to a fluid coking zone, maintaining said solids therein at a temperature between about 950 and 1600 F. while passing a stripping gas upwardly therethrough, recovering vapors rich in hydrogen and methane from said coking zone, withdrawing dried solids from said coking zone, and reheating and recycling at least a portion of said withdrawn dried solids to the process.

3. A process according to claim 2 wherein the contact time of the vapors with the hot solids in said conversion zone is limited to about 0.01 to 1.5 seconds.

4. A process according to claim 2 wherein the solids separated from the conversion zone product are passed to said coking zone.

5. A process according to claim 4 wherein a portion w ma of the solids withdrawn from the coking zone is introduced into the vaporizing zone, and another portion of the lastmentioned solids is passed to a combustion zone where these solids are heated to at least 1300 F. by combustion with an oxygen-containing gas, and the reheated solids are passed from the combustion zone to said conversion zone.

6. A process for converting a heavy petroleum residuum into olefinic and aromatic hydrocarbons which comprises mixing said residuum with finely divided coke particles, passing the resulting mixture in the form of a gasiform suspension through a confined elongated vaporizing zone at a temperature between about 950 and 1100 F. until said residuum is partially but not completely vaporized, separating the resulting vapors from the coke particles and unvaporized liquid hydrocarbons adhering thereto, passing the separated wet coke particles to a fluid coking zone, passing a stripping gas upwardly through the particles in the coking zone until the liquid hydrocarbons are converted into a solid residue and gases rich in hydrogen and methane, withdrawing said gases from the coking zone, withdrawing dried coke particles from said coking zone, mixing the aforesaid separated hydrocarbon vapors with dry finely divided coke particles, passing the resulting mixture in the form of a gasiform suspension through a confined elongated conversion zone at a temperature between about 1300 and 1600 F. and at a velocity corresponding to a vapor residence time therein of about 0.01 to 1.5 seconds, separating the resulting cracked vapors from the high temperature solids, quenching the separated cracked vapors to a temperature below about 600 F., and passing the separated hightemperature solids to the vaporizing zone for mixing with additional residuum feed.

7. A process according to claim 6 wherein the said conversion zone comprises a plurality of distinct elongated reaction zones through which the hydrocarbon vapors pass in series, the hottest solids being introduced to the inlet of the last reaction zone of the series, and the solids separated at the outlet of each reaction zone being introduced to the inlet of the next preceding zone.

87 In a process wherein olefinic and aromatic hydrocarbons are produced from a heavy residual hydrocarbon feed boiling at least in substantial part above about 950 F. by contacting said feed with hot substantially inert powdered contact solids, the improvement which comprises mixing said feed with the hot contact solids so as to produce a mixture heated to about 950 to 1100 F., passing this mixture as a dilute gasiform suspension through a confined elongated vaporizing zone at a high velocity corresponding to a residence time of about 0.3 to 8 seconds in said vaporizing zone, separating the resulting hydrocarbon vapors from the suspended wet solids which contain incompletely vaporized feed adhering thereto, mixing the separated hydrocarbon Vapors with contact solids heated to a temperature at least 100 F. higher than the aforesaid solids introduced into said vaporizing zone, passing the resulting mixture as a dilute gasiform suspension through an intermediate confined elongated conversion zone at a velocity corresponding to a residence time of about 0.01 to 1.5 seconds therein, separating the resulting cracked hydrocarbon vapors from the 10 solids suspended therein, mixing the separated cracked hydrocarbon vapors with contact solids in ratio of about 10 to parts of solids per part of hydrocarbon, said solids being heated to a temperature of at least 1300 F. and at least 100 F. higher than the aforesaid solids introduced into the said intermediate confined elongated con version zone, passing the resulting suspension in the form of a gasiform suspension through a last confined elongated conversion zone at a temperature between about 1300 to 1600 F. and a residence time of about 0.01 to 1.5 seconds, separating the resulting cracked product vapors from the solids suspended therein, quenching the separated product vapors by mixing with an inert cooling medium, recovering the quenched products, passing the solids separated after passage through a given elongated zone to the inlet of the next preceding elongated zone, the wet solids separated from the first named vaporizing zone being passed to a fluid coking zone where the said solids are dried, and passing the dried solids from the coking zone to a combustion zone where the solids are heated to a temperature sufiicient for feeding to the aforesaid last conversion zone.

9. An apparatus for converting heavy hydrocarbon feed into olefins, which comprises a substantially vertical elongated vaporizer conduit, means for feeding said hydrocarbon feed and finely subdivided solids into the lower end of said vaporizer conduit, separating means for separating vapors from the powdered solids at the upper end of said vaporizer conduit, a substantially vertical clongated conversion conduit communicating at its bottom inlet end both with the vapor outlet of the aforesaid separating means and with a reservoir of hot dry powdered solids, which reservoir is separate and independent from said first-named means for feeding solids other separating means at the outlet end of the said conversion conduit adapted to separate vapors from solids, and means for passing solids from said last-named separating means to said first-named means for feeding solids.

10. An apparatus according to claim 9 further comprising a coking chamber adapted to receive the solids separated from the said vaporizer conduit, means for injecting a fiuidizing gas upwardly through said chamber, means for removing gases from an upper part of the chamber, and means for removing dried powdered solids from said chamber.

11. An apparatus according to claim 9 wherein both said separating means consist essentially of the terminal portion of the respective conduit arranged in the shape of a U-bend and adapted for removal of substantially solidfree vapors from the inner wall of said U-bend.

References Cited in the file of this patent UNITED STATES PATENTS 2,378,531 Becker June 19, 1945 2,396,109 Martin Mar. 5, 1946 2,526,696 Schutte Oct. 24, 1950 2,543,884 Weikart Mar. 6, 1951 2,608,526 Rex Aug. 26, 1952 2,687,992 Leffer Aug. 31, 1954 2,718,491 Green Sept. 20, 1955 2,731,508 Iahnig et al. Jan. 17, 1956 2,763,600 Adams et al Sept. 18, 1956 

1. A PROCESS FOR CONVERTING HEAVY OILS CONTAINING CONSTITUENTS NONVAPORIZABLE WITHOUT CRACKING WHICH COM-PRISES THE STEPS OF VAPORIZING A HEAVY OIL IN AN ELONGATED VAPORIZATION ZONE BY PASSING SAID HEAVY OIL AND HEATED PARTICULATE SOLIDS THERE THROUGH AT A TEMPERATURE BETWEEN ABOUT 950* AND 1100*F., SEPARATING TARRY SOLIDS CONTAINING ADHERING LIQUID HYDROCARBONS FROM THE VAPORS SO FORMED, CONTACTING SAID VAPORS WITH ADDITIONAL AMOUNTS OF PARTICULATE SOLIDS AT A CHEMICALS COKING TEMPERATURE AND PASSING THE RESULTING MIXTURE THROUGH AN ELONGATED CONVERSION ZONE, SEPARATING FROM THE EFFLUENT OF SAID ELONGATED CONVERSION ZONE GASIFORM CONVERSION PRODUCTS AND SPENT SOLIDS, PASSING SAID SPENT SOLIDS AND SAID TARRY SOLIDS TO A FLUID COKING ZONE AND COMPLETELY DRYING SAID SOLIDS THEREIN, RECOVERING GASES COMPRISING HYDROGEN AND METHANE FROM SAID FLUID COKING ZONE, PASSING A PORTION OF THE DRIED SOLIDS FROM SAID FLUID COKING ZONE TO SAID ELONGATED VAPORIZATION ZONE, HEATING THE REMAINDER OF THE DRIED SOLIDS TO ABOVE SAID CHEMICALS COKING TEMPERATURE AND TRANSFERRING THE SOLIDS SO HEATED TO SAID ELONGATED CONVERSION ZONE. 